Ethylene is one of the most dominant base petrochemicals, with a global production rate of approximately 180 million metric tons per annum production. Over sixty percent is used in the production of polyethylene with different properties such as high density polyethylene (HDPE), low density polyethylene (LDPE), and linear low-density polyethylene (LLDPE). Ethylene may be produced by steam cracking of hydrocarbons, predominantly saturated hydrocarbons. In steam cracking, gaseous feedstocks such as ethane, ethane and propane, and/or a mixture of propane and butane in the form of a liquefied petroleum gas (LPG), may be fed to a cracking furnace. Cracking furnaces operating on gaseous hydrocarbon feedstocks may be referred to as a gas cracker and cracking furnaces operating on liquid hydrocarbon feedstocks such as naphthas, and/or gas oils, may be referred to as liquid crackers. In general, the hydrocarbon feedstock to the steam cracker may be diluted with steam and thereafter briefly be exposed to a high temperature environment within a cracking furnace to produce a product gas comprising ethylene. Typically, the reaction temperature is relatively high, at approximately 850° C., but the reaction is only preferred to take place with minimal residence time. After the cracking temperature has been reached, the product gas may be quickly quenched to stop the reaction, for example, by quenching the product gas in a transfer line heat exchanger.
The product gas produced in the steam cracker may depend on the composition of the feedstock, the steam-to-hydrocarbon weight ratio, cracking temperature, and reaction residence time. Light hydrocarbon feeds such as ethane, LPG, or light naphtha may result in cracked product streams rich in lower olefins such as ethylene, propylene, and mixed butylenes. Liquid hydrocarbon feeds may yield the same lower olefins but additionally may produce aromatic-rich hydrocarbons and hydrocarbons which may be suitable for inclusion in gasoline or fuel oil.
A higher cracking temperature, also referred to as severity, may favor the production of ethylene and benzene, whereas lower severity may produce higher amounts of propylene, mixed butylenes and liquid products. The thermal process within the cracking furnaces may also result in the gradual deposition of coke on the inner walls of the radiant coils. The deposition may degrade the efficiency of the cracking furnaces, so design features of the radiant coils have evolved to minimize coke formation. Nonetheless, a cracking furnace can usually only maintain operation for a few months at a time between de-coke cycles. Decoking requires the furnace to be isolated from the process and then a flow of steam followed by a steam/air mixture followed by air-only through the radiant coils. This converts the hard solid carbon layer to carbon monoxide and carbon dioxide. Once this reaction is complete, the cracking furnace can be returned to service.
Steam crackers may be designed and built to favor ethylene production. As previously discussed, other byproducts may also be produced from steam cracking. For example, the co-production of these byproducts such as propylene, butanes and butenes, and aromatic pyrolysis gasoline may be dependent on the steam cracker design and feed composition, as illustrated in Table 1. Table 1 illustrates an example byproduct composition based on production of 100 tons of ethylene for different feedstocks. AGO is atmospheric gas oil.
TABLE 1FeedstockEthanePropanen-C4/i-C4NaphthaAGOProductH2 + CO7.744.293.883.322.86CH4758.9468.5349.4843.02C2H20.951.691.582.221.37C2H4100100100100100C2H666.2711.5311.5611.0311.07C3H62.1933.4962.5452.2657.46C3H80.2318.712.171.231.25C4's4.2710.1640.6430.7538.67Pyrolysis1.6610.7720.262.3182.9GasolinePyrolysis0.312.83.979.763.5Fuel Oil
As one of ordinary skill in the art will appreciate, byproduct yields are also affected by the design of the cracking furnace and the severity of cracking. Acetylene is a byproduct of steam cracking. The yield of acetylene may be about 0.95 to about 2.22 tons per 100 tons of ethylene, depending on the feedstock and cracking severity. Light naphtha is typically cracked at the highest severity and may produce the most acetylene. Conversely, gas oils may be cracked at a relatively lower severity to reduce the rate of coking in the radiant coils. Consequently, the production of acetylene from atmospheric gas oils may be relatively lower than alternate feedstocks as shown in Table 1. There may be a significant range acetylene produced from a given steam cracker as feedstocks are varied. In general, the acetylene content of a marketable ethylene product must be lower than about 1 parts per million by weight (ppm). Acetylene removal is a critical step in the manufacture of ethylene, since failure to satisfy the acetylene content in the final ethylene specification may result in an unsaleable product.
Many different techniques are utilized in the removal of acetylene in the purification of product ethylene. This technology has progressed through distinct stages and encompasses various methods.
A method of acetylene removal may be solvent extraction. In the solvent extraction method, the acetylene is selectively absorbed into a solvent, such as n,n-dimethyl formamide (DMF), followed by its purification resulting in a merchantable acetylene product.
Another method of acetylene removal may be selective acetylene hydrogenation. The reactions that can occur in over the catalyst bed in acetylene hydrogenation units are listed below. An acetylene hydrogenation chemistry map is illustrated in FIG. 1.C2H2+H2→C2H4 (desirable, causes ethylene gain)C2H4+H2→C2H6 (undesirable)2C2H2+H2→C4H6 (undesirable)C4H6+H2→C4H8 (undesirable)nC2H2+nH2→Green Oil (undesirable oligomer, causes fouling)
Green oil formation and fouling may be a major issue in back-end hydrogenation reactors. The acetylene hydrogenation reactor may be placed at various different locations in the ethylene plant process flowsheet. For example, the hydrogenation reactor may treat a raw gas stream wherein the reactor may be placed downstream of a caustic tower. Another example may be a front-end selective catalytic hydrogenation reactor placed upstream of a demethanizer (DC1) treating either a front-end deethanizer (DC2) overhead stream, or a front-end depropanizer (DC3) overhead stream. Another example may be back-end catalytic hydrogenation reactors placed downstream of the DC1 treating the DC2 overhead stream. Some suitable catalysts may include, without limitation, nickel based catalyst for raw gas acetylene hydrogenation reactors, and palladium and/or bimetallic palladium/silver catalysts for front-end and back-end acetylene hydrogenation reactors.
For a raw cracked gas selective hydrogenation application, the reactor may be placed within the process gas compressor circuit, downstream of the caustic tower, but upstream of the main driers. Hence, the feed to this acetylene hydrogenation unit may be wet C4's and lighter, containing some C5's. Historically, this unit has been designed for operation with sulfided Ni/Co/Cr catalysts. Such acetylene hydrogenation units have been installed in many gas (ethane-only) crackers designed and built in the 1960's/early 70's and some remain in operation today. This operation may result in hydrogenation of C4 acetylenes and butadiene, hence minimizing the potential for fouling the bottoms section of downstream DC2 and/or DC3. Additionally, it may convert the majority of C3 and C4 acetylenes and diolefins into olefins, making the final C3 and C4 streams more attractive for fuel gas, and/or for recycle cracking. In such a process, the catalyst deactivation may be relatively rapid, especially for the first bed. This is primarily due to the presence of C4 and C5+ acetylenic and diolefinic species. Hence, a spare reactor system is needed for continuous operation of 5+ years. Typically, this hydrogenation unit operates with a net ethylene loss.
For a front-end DC2 and front-end DC3 selective hydrogenation process, the acetylene reactors may precede the DC1 in the process flow scheme. As a result, these reactor feeds contain a large excess of hydrogen, typically 10 to 35 mol percent. In a front-end DC2 configuration, the DC2 is the first distillation column and the reactors are on the overhead stream. Thus, the feed contains a C2 and lighter components. Similarly for a front-end DC3 configuration, the initial distillation column is the DC3. As the acetylene reactors are on the overhead of this column, the gaseous feed to the reactors is composed of C3 and lighter hydrocarbon. The feed to the front-end selective hydrogenation reactors may be characterized as a clean stream as the high concentration of H2 may suppress green oil formation. Thus, the design does need a spare reactor. The high concentration of H2 may result in the operation being sensitive to initiation of ethylene hydrogenation and this may result in exothermic runaways and poor operational stability due to sensitive to CO fluctuations.
In back-end selective hydrogenation reactor, the acetylene is typically contained in a C2-rich stream whereby stoichiometric amounts of hydrogen, and in some cases small amounts of carbon monoxide, are added to control the extent of acetylene hydrogenation. In this type of application, the reactors may be located at the overhead of the DC2. This hydrogenation configuration may be suitable for moderate acetylene concentrations when the recovery of acetylene is not of interest. Hydrogen may be added in a molar ratio of 1.2 to 2.2 relative to acetylene in the DC2 overhead gas and the mixture may be passed over a fixed bed palladium-based catalyst. The hydrogenated vapor from the reactor system flows through the guard dryer (often referred to as secondary dehydrator) to the ethylene-ethane fractionator (ethylene splitter). The reactor effluent typically contains less than 1 ppm of acetylene but is contaminated with traces of hydrogen and methane which represent the major disadvantage of a back-end acetylene hydrogenation system. The unreacted H2 in the reactor effluent must be removed via a pasteurization section at the top of the ethylene/ethane splitter or via a downstream secondary DC1.
Green Oil, which may be oligomers of acetylene, may cause fairly rapid catalyst deactivation. Therefore, it may be necessary to include spare reactors with facilities for in-situ catalyst regeneration and reduction. Each reactor may be regenerated about 1-4 times a year. Back-end units may have downstream green oil (C6+) removal facilities. Yet this location has the advantage over the front-end hydrogenation in that it allows very accurate control of hydrogen concentration and reaction temperature, which may result in a higher selectivity of the reaction and higher ethylene gain. Green oil polymer may be formed by side reactions of the hydrogenation of acetylene to ethylene and ethane over Pd, Ni, Pt, etc. catalyst. Green oil may occur due the dimerization of acetylene in the presence of hydrogen to butadiene followed by oligomerization with successive addition of acetylene to a chain of molecules adsorbed on the catalyst surface. The amount of green oil formed may decrease as the hydrogen partial pressure increases. The green oil may be a mixture mainly C4 to C20 reactive oligomers of varying composition. The heavier fraction may be adsorbed on the pores catalyst causing eventual loss of the catalyst activity and thus requiring regeneration by steaming out the deposited green oil. The light end components of the green oil remain in the gas phase, part of which may condense into fine droplets with the gas stream leaving the reactor. These fine droplets may cause fouling of the downstream equipment.
Catalyst deactivation by green oil may become a major problem at very low hydrogen to acetylene ratios. Green oil formation may be decreased by the use of silver promoted Pd catalyst on Al2O3 which help terminate the chain growth at the butadiene stage. Thus, instead of the formation of heavier diolefins oligomers, butadiene exits with the gas. This new catalyst generally reduces the formation of green oil to third or half the amount formed with the non-promoted catalysts. The concentration of the green oil in the gas leaving the hydrogenation reactor is in the order of about 100 ppm to about 1000 ppm dependent upon the operating temperature, age of the catalyst, CO content of the gas, H2/acetylene ratio, etc. The droplet size of green oil condensing in the gas stream downstream of the reactor may be mostly less than 5 micron size. The hydrogenation units may be designed with multi-stage adiabatic reactors with inter-stage cooling to remove the heat generated by the exothermic hydrogenation reactions of acetylene.
The amount of green oil formed may be primarily function of the concentration of acetylene being converted. Hence, the rate of formation of green oil may be higher from the first bed. Green oil formation may decrease with increasing the partial pressure of H2, which is the main reason for the much lower green oil formed in front-end hydrogenation system as compared to back-end hydrogenation units. Typically in back-end hydrogenation reactors 10-20% of the acetylene is converted to C4 and heavier green oil.
Ethylene plants are expected to operate for a period of 5-7 years between turnarounds. Hence, a spare bed is provided to allow the plant to continue operating, when one of the bed(s) is fouled. The fouled bed is taken out of operation, and the spare bed is placed in service. The green oil is then drained and the fouled bed is regenerated and put on standby mode.
The gas leaving the hydrogenation reactor may be cooled, and more green oil may condense from the vapor phase into fine droplets, which may deposit on downstream heat exchangers, dehydrator beds, and on ethylene fractionator internals. The depositing droplets are polymeric and cause fouling of the equipment thus eventually requiring expensive unplanned shutdowns to clean-up the deposited green oil.
Fuel gas used for the regeneration of the secondary dehydrators may strip out the deposited green oil on the molecular sieves, making the fuel gas contaminated with the green oil. The contaminated gas may cause fouling of the low NOx burner nozzles which may lead to lower furnace efficiency and more frequent and costly burner tip cleaning.
Back-end acetylene hydrogenation system may be characterized by substantial formation of green oil and the need to include green oil removal system to protect the downstream secondary dehydrator and ethylene purification equipment. Different industrial methods may be used for the separation of green oil from the hydrogenation reactor gaseous effluent stream were evaluated including: washing of the wet gas stream from the reactor with a liquid ethylene stream in an absorption tower which may be the most efficient method capable of removing>99.5% of the green oil droplets. Impaction of the wet gas through a packed bed wherein the green oil removal may be sensitive to gas flow rate and distribution through the bed. Alternative impaction separation may be accomplished by a mesh pad in a knock-out drum. However, this may be the least efficient removing less than 70% of the green oil droplets.